Two stage hydrocarbon conversion process with hydrocracking in both stages to produce a high octane gasoline



8 NN M 1. NN S S 8G s Y 0N RU 8m OA E 1C .....moow TN.. N 3am N N MM EH m mw u.. wm T A Il HG d U WA HE V L HM N. m WT WN V I. sc QN N502 O M Ew% N Smml SP N UNAS O t Mmmw D mmh/11:55 qe... \2 Cmw e. .u oOwwIoOOm W Hmmm S a W .C d WNwm w w E ll. 2 ooww mm u mzoN n Nv NN mm m u uzooome: mroomof rz OS I.. R O um. N wm A N n Mw l m W/ SMN mm. u... 9 n nowl U .Va %T1 I w w u mzoN 1M |U QM. m L. m 02U OOD I Dumm ,w o 7T 1 w m u. M.. m .4. .n N D. mw A v E b 7 I U NT- INTRODUCTION This invention relates to a hydrocarbon conversion process, more particularly to -a hydrocarbon conversion process for converting petroleum distillates and residua into various valuable products, and still more particularly to a catalytic conversion process capable of producing middle distillates, light gasoline and isobutane.

PRIOR ART HYDROCRACKING OF HYDROCAR- A BON FEEDS TO PRODUCE MIDDLE DISTIL- LATES AND GASOLINE, AND PROBLEMS IN- VOLVED A. Nitrogen content of feed-It is Well known that nitrogen in a hydrocarbon feed is deleterious to certain hydrocracking catalysts, particularly highly acidic hydrocraclcing catalysts, and that, in order to provide a practical process for producing middle distillates from a feed containing substantial amounts of nitrogen, a catalyst having no more than weak acidity has been necessary so that the deleterious eifect of nitrogen on the catalyst would be minimized. However, catalysts having no more than weak acidity frequently result in thev production of substantial quantities of heavy gasoline, which often is a less desirable product than middle distillates, light gasoline and isobutane.

B. Ratio of s-C4 to normal-C4 product-It is well known that a high iso-C., to normal-C4 product from a liydrocracking zone is highly desirable. Isobutane, for example, is a valuable product for use in motor gasoline blending, Whereas normal butane is less valuable. A low isoC4 to normal-C4 product ratio has been a disadvantage of many prior art processes.

C. Parfznicy of unconverted bottoms fraction recycled to reactorf-It is well known that various prior art catalysts, particularly catalysts of extremely low activity which are useful in the production of middle distillates and catalysts of extremely high activity which are useful in the production of gasoline, produce an unconverted bottoms fraction having a high content of normal paraiiins. It is known that these normal parans are deleterious -to the hydrocracking operation because they are extremely refractory to further hydrocracking and therefore, particularly Where high middle distillate production is desired, as a practical matter, cannot be recycled. It would be desirable if a two-stage process were e United States Patent ice sirable if the catalysts used in processes for meeting the foregoing prior art problems were regenerable.

OBJECTS vide such a process capable of producing large quantities of isobutane.

It is a further object of the present invention to provide such a process and catalysts wherein the unconverted bottoms fraction from the trst stage has a sufficiently low content of normal parains to permit recycling a selected portion of this fraction to the rst stage reactor in sustained recycle operation, to increase the overall yield of light gasoline and isobutane.

It is a further object of the present invention to provide such a process wherein the catalysts in each stage may be regenerated to reimpart to them a substantial portion of their original fresh activity.

DRAWING Theinvention will best be understood, and further objects and advantages thereof will be apparent,'from the following description when read in conjunction with the accompanying drawing which is a diagrammatic illustration of process units and flow paths suitable for carrying out the process of the present invention.

STATEMENT -OF INVENTION In accordance with the present invention, there is provided a process for producing gasoline and middle distillates from a hydrocarbon feed selected from the group consistingof petroleum distillates boiling from 200 to 1100J F. and petroleum residua boiling above 500 F., which comprises contacting said feed and from 1000 to 10,000 s.`c.f. of hydrogen per barrel of said feed in a first conversion zone at a temperature from 500 to 950 F., a pressure above 500 p.s.i.g. and an LHSV of 0.1 to 4.0 with a catalyst comprising at least one hydrogenating component selected fromthe Group VI metals and compounds ofGroup VI metals and at least one hydrogenating component selected from the Group VIII metals and compounds of Group VIII metals intimately yassociated with a catalyst support comprising silica-magnesia, `re-l covering from the eiiiuent from said iirst conversion zone available for producing middle distillates, light gasoline i and isobutane in a rst stage and light gasoline and isof butane 'in a second stage without producing a irst stage a middle distillate product and a light gasoline pro`duct hydrocracking in a second conversion zone in the presence of an active acidic hydrocracking catalyst at a pressure of 500 to 3000 p.s.i.g; and a temperature of 550 to 850 F. a portion of the elucnt from said rst conversion zone boiling from about 180 to 300 F., to produce a light. gasoline product having an increased isoparaffin content and recovering said light gasoline product from said second conversion zone. i

Further in accordance with the present invention, there is provided a process for producing gasoline and middle distillates which comprises converting a heavy gas oil feed in a first conversion zone in the presence of a catalyst comprising nickel and tungsten on a silica-magnesia support at a temperature of from 500 to 950 F., a pressure above 500 p.s.i.g. and an LHSV of from 0.4 to 2.0, to a total product containing above 60 volume percent, based on feed converted, of 320 to 650 F. middle distillate, separating said total product into fractions including a middle distillate fraction and a gasoline fraction, recovering said middle distillate and gasoline fractions as products, hydrocracking in a second conversion zone in the presence of an active acidic hydrocracking catalyst at a pressure of 500 to 3000 p.s.i.g. and a temperature of 500 to 850 F. a portion of the effluent from said first conversion zone boiling from about 180 to 300 F., recovering a gasoline product from said second conversion zone, and recycling to said first conversion zone at least a portion of the effluent from said first conversion zone boiling above about 660 F.

Still further in accordance with the present invention, there is provided a process for converting a nitrogen-containing hydrocarbon feed selected from the group consisting of petroleum distillates boiling from 500 to ll00 F. and petroleum residua boiling above v500" F. which comprises concurrently hydrofining and hydrocracking said feed by contacting with feed in a first stage in the presence of from 1000 to 10,000 s.c.f. of hydrogen per barrel of said feed and in the presence of a catalyst comprising at least one hydrogenating component selected from the group consisting of Group VI metals and compounds thereof and at least one hydrogenating component selected from the group consisting of Group VIII metals and compounds thereof intimately associated with a silicamagnesia catalyst support at a temperature of 500 to 950 F., a pressure above 500 p.s.i.g. and an LI-ISV of from 0.1 to 4.0, recovering a gasoline product and a middle distillate product, hydrocracking in a second stage in the presence of an active acidic hydrocracking catalyst at a pressure of 500 to 3000 p.s.i.g. and a temperature of 550 to 850 F. at least a substantial portion of the liquid effluent from said first stage, recovering a gasoline product from said second stage, and recycling to said first stage at least a portion of the effluent from said first stage boiling above about 660 F.

Still further in accordance with the present invention, there is provided a process as aforesaid wherein at least two reactors are used in said first conversion zone, each containing said silica-magnesia supported catalyst, and wherein said reactors are so arranged that they can be switched from parallel, for maximizing middle distillate i HYDROCARBON FEEDS SUITABLE FOR USE IN THE PROCESS OF THE PRESENT INVENTION Suitable feeds for use in the process of the present invention are petroleum distillates boiling from 200 to 1100 F., preferably petroleum distillates boiling from 500 to l100 F., and petroleum residua boiling above 500 F., and mixtures of the foregoing. Heavy gas oils and catalytic cycle oils are excellent feeds to the process as well as conventional FCC feeds and portions thereof. Residual feeds may include Minas and other paraffinictype residua.

Particularly when it is desired to produce middle distillates, including jet fuels, which are exceptionally high in naphthene content and low in aromatic content (therefore having high smoke points) and low in normal paraffin content (therefore having low freeze points), it is preferable to use a feed in the process of the present invention which has an initial boiling point of 500 F. or above. Where the feed has an initial boiling point above 500 F., it is converted in the process of the present invention did rectly to a synthetic material, i.e., one boiling below the initial boiling point of the feed, which is a preferred jet fuel or middle distillate having high naphthene content, low normal paraffin content and therefore low freeze point, and low aromatic content and therefore exceptionally high smoke point. It has been found that feeds having lower initial boiling points, for example around 300 to 400 F., tend to produce excessive quantities of nonsynthetic products having a high aromatics contents and therefore exceptionally low smoke points, although the freeze point may be satisfactory. Such a nonsynthetic product also tends to have a high pour point.

FIRST CONVERSION ZONE IN PROCESS OF PRES- ENT INVENTION, AND NITROGEN CONTENT OF FEED THERETO It has been found that the silica-magnesia supported hydrocracking catalyst in the first conversion zone of the process of the present invention is relatively nitrogen insensitive, compared with conventional acidic hydrocracking catalysts such as nickel sulfide on silica-alumina. Accordingly, the nitrogen content of the feed used in the process of the present invention may be relatively high, and excellent hydrocracking results still may be obtained in said conversion zone at reasonable temperatures, without the necessity for rapidly raising the temperature to maintain conversion as is necessary when hydrocracking a high nitrogen content feed over a conventional acidic hydrocracking catalyst such as nickel sulfide on silicaalumina. Although high nitrogen content feeds can be tolerated by said first conversion zone hydrocracking catalyst, it will be noted that said catalyst also is an excellent hydrodenitrification catalyst, and is efficient in concurrently hydrofining as well as in hydrocrackng the feed. Nevertheless, the process of the present invention may be rendered even more efiicient if the feed either is low in nitrogen content or first is hydrofined by conventional methods prior to being hydrocracked in said first conversion zone in accordance with the process of the present invention. And in certain applications a conventional hydrofining zone following said rst conversion zone is desirable; as will be discussed below, in one embodiment of the present invention, wherein very heavy feeds, for example propane deasphalted residua, are used, the feed may be processed in three stages; in the first stage, the feed may be concurrently hydrocracked and denitrified to a large extent over a silica-magnesia supported catalyst, following which a portion of the effluent from the first stage may be further denitrified in a second stage before being hydrocracked over an acidic hydrocracking catalyst in a third stage.

Generally speaking, it is possible to operate the first conversion Zone in the process of the present invention at slightly lower temperatures when the feed has a low nitrogen content, for example, from 0 to 10 p.p.m. total nitrogen, than temperatures that are necessary for the same conversion when the feed has a high nitrogen content, for example from 10 to 1000 p.p.m. total nitrogen. However, even feeds containing considerably higher levels of nitrogen than 1000 ppm. total nitrogen may be satisfactorily converted in the process of the present invention to valuable products, contrary to conventional prior art processes wherein acidic hydrocracking catalysts, such as nickel sulfide on silica-alumina, are used. In such conventional processes, it is impossible as a practical matter to use feeds with such high nitrogen contents.

The catalyst in the first conversion zone in the process of the present invention is capable of concurrently accomplishing both denitrification and hydrocracking. The hydrocracking facilitates the concurrent denitrification because, upon the breaking of carbon-to-carbon bonds, nitrogen is more easily removed. At higher levels of cracking conversion, somewhat higher pressures may be desired to counteract catalyst fouling and deactivation.

The nitrogen compounds tend to concentrate in the arsenic the other one of the two groups. Further information regarding the denitritication activity of the catalyst is set forth in Table V below.

G. Selectviy of catalyst for middle distillate production-The catalyst of the present invention has a high selectivity for the production of middle distillates from various hydrocarbon feeds. It has a much greater selectivity for the production of middle distillates than conventional acidic hydrocracking catalysts, such as nickel sulfide on silica-alumina. The high yields of rniddie distillates resulting from the selectivity of the catalyst of the present invention for middle distillate product is unexpected in View of the selectivity for gasoline production that is characteristic of many prior art hydrocracking catalysts, for example nickel suliide on silicaalumina. Further information regarding the selectivity of the present invention catalyst for the production of middle distillates is set forth in Table III below.

DESCRIPTION OF PROCESS FLOW ARRANGE- MENTS SUITABLE FOR CARRYING OUT THE PROCESS OF THE FRESENT INVENTION Referring now to the drawing, there shown is a diagrammatic illustration of an embodiment of process units and flow paths suitable for carrying out the process of the present invention.

A hydrocarbon feed is passe-d through line l into contact in hydrocracking zone 2 With the aforesaid silicamagnesia supported catalyst and with hydrogen entering zone 2 through line 3 under the hydrocracking conditions previously discussed. From zone 2 an effluent is passed through line 4 to separation zone S from which hydrogen is recycled through line 6, ammonia is withdrawn through line 7, and remaining materials are passed through line 8 to separation zone 9. From separation zone 9 a C4 and lighter stream, including isobutane, is Withdrawn through line 10, and remaining materials are passed through line 1S to separation zone 16. From separation zone 15, a light gasoline product is Withdrawn through line 17, a heavy gasoline boiling from `about 180 to 300 F. is passed through line 10 to hydrocracking zone 19, and materials heavier than about 300 F. are passed through line to separation zone 25.

From separation zone 2.5 middle distillate products boiling from about 300 to 660 F. are withdrawn through line 2o and materials boiling above about 660 F. are recycled through line 27 to hydrocracking zone 2. If desired, a minor portion of the materials in line 27 may be Withdrawn from the system through line 28. If desired the hydrocarbon feed to hydrocracking zone i9, entering that zone through line 18, may be augmented by additional hydrocarbon stocks boiling between about 180 and 300 F., for example straight run stocks, through line 29. Hydrocracking zone 19 may contain a conventional acidic hydrocracking catalyst, for example nickel suliide or silica-alumina, platinum on silica-alumina, etc., and may operate under conventional hydrocracking conditions, for example a pressure of from 500 to 3000 p.s.i.g. and a temperature of from 550 to 850 F. lt is well known that such catalysts can be subjected to regeneration with -an oxygen-containing gas under conventional regeneration conditions. Hydrocracking zone 19 is supplied with hydrogen through line 35. Zone 19 effluent is passed through line 36 to separation zone 37. From separation zone 37, C4 and lighter hydrocarbon gases, including isobutane, are Withdrawn through line 38, and a C5 :plus light gasoline product is withdrawn through line 39.

Hydrocracking zone 2 may comprise two hydrocracking reactors, each containing the catalyst of the present invention and each operating under the aforesaid process conditions. These two reactors may be arranged in a known manner so that alternately they can be connected in parallel and in series. When connected in parallel,

they will operate to maximize middle distillate production, and when switched to series operation they may maximize gasoline or middle distillate production. In series operation, middle distillate production may be maximized by withdrawing middle distillate as a product from the first reactor as well as from the second, for example from an inter-reactor fractionation zone. In series operation, gasoline production may be maximized by including the middle distillate produced in the first reactor in the feed to the second reactor. In either series arrangement, it is preferred to remove from the system any ammonia produced in the first reactor, rather than permitting it to pass to the second reactor. Such switching arrangements will enable the ratio of middle distillate to gasoline product to be varied in order to achieve further process application flexibility. In series operation to produce gasoline, where ammonia formed in the rst reactor has been removed, the second reactor, because it is operating with a feed that has been denitriiied in the irst reactor, is operable at lower temperatures, thus providing leeway for increase in severity of the operating conditions in the second reactor to increase gasoline production. The resulting gasoline, produced over the catalyst of the present invention, is isoparatiinic and of high quality, in contrast to the normal parainic character of gasoline produced over hydrocracking catalysts having weak acidity.

Because the catalyst in zone 2 serves as an effective hydroiining catalyst, the materials in line 8 are low in nitrogen and therefore are specially suitable for further hydrocracking in the presence of the acidic catalyst in zone 19.

The process of the present invention is specially effe..- tive for converting heavy feed such as residua and propane deasphalted oils when a conventional denitriication zone is inserted between the first hydrocracking zone and the second hydrocracking zone of the process. Because such feeds generally are specially difcult to denitrify, and because for most eliicient results the feed to the second conversion zone 19 here, containing an acidic hydrocracking catalyst, should have a minimum nitrogen level, the insertion of a conventional denitriication zone between the two hydrocracking zones in the present process can be of significant value.

The conventional denitriiication zone may be inserted for example in line 8 or line 15, and may be operated under conventional denitriiication conditions with either the silica-magnesia supported catalyst used in hydrocracking zone 2, or with any conventional denitriication catalyst. Such a three-stage process enables the heavier feeds to be hydrocracked and partially denitriiied in the first stage, thereby reducing both the molecular weight and the nitrogen level of the feed and greatly accelerating the rate of the remaining denitriiication to be accomplished in the second, or conventional, denitriiication zone.

TABLE L Conranrson on FIRST STAGE cA'raLrsr 0F PROCESS OF PRESENT INVENTION WITH CON- VENTIONAL CATALYSTS RE STARTING TEMPERA- TURES AND FOULING RATES The following table sets forth on a comparative basis single stage hydrocracking results of processing a 650 to 980 F. heavy Arabian gas oil having a total nitrogen content of 660 to 700 p.p.m. at the indicated average catalyst temperature, about 50 to 55 volume percent substantially constant per-pass conversion to products boiling below the initial boiling point of the feed, 1.0 LHSV, 2000 p.s.i.g. and a hydrogen rate sufficient to permit withdrawal from the hydrocracking zone of 4500 s.c.f. of hydrogen per barrel of feed, over the first stage catalyst of the present invention compared with hydrocracking the same feed under the same conditions over various prior art catalysts. The factors compared are: (l) the average catalyst temperature necessary to givc said substantially constant 50 to 55% per-pass conversion, which substantially constant conversion is indicated by the substantially 9. constant product gravity shown; and (2) the catalyst fouling rate.

o From the above table it will be noted that: (1) as acidity increases, the product iso to normal ratio increases Support Hydrogenating component, Av. eat.

y percent Area, temp., F. Product Cat. No. M2/g. necessary gravity Fouling rate Y for desired SiO2-Al203 SiO2-Mg0 Ni W Mu Pt conversion 27% MgO. 7. 0 19. 3 316 759 40. 0 None observab1e.1

27% MgO... 4. 9 22. 8 254 '755 40, 3 D0.l 27% MgO Y 5.0 5. 0 756 39. 9 Moderate.2 27% MgO.-. 5.0 1. 8 767 39. 5 DOJ! 27% MgO 8. 5 437 790 39. 5 High.3 27% MgO 9. 0 445 765 40. 3 Very high .4 27% Mg 0.5 845 38.5 Do.4

A120 4. 5 12. 2 134 790 40. 3 None observable.1 28% A12O2 3. 9 10. 4 130 792 40.0 Moderate.2 47% A1203... 5.0 7. 0 780 40.0 Do.2 47% A1203." 4. 4 9. 4 92 '790 39. 8 D0.2 10% A1203--. 18. 0 116 805 39. 8 .3

1 0.05 F. per hour. 2 0.10-0.15 F. per hour.

From the above table, it will be noted that only catalysts 1 to 4 resulted in both (l) the desired conversion rate at a reasonably low average catalyst temperature, in each case 767 F. or below, and (2) a reasonably low catalyst fouling rate, in each case, moderate, as defined, or less. It will be noted that catalysts 5 to 7, each hav-v ing one hydrogenating component only, on a silica-magnesia support, resulted in an excessive catalyst fouling rate, i.e., one that was high, as defined, orhigher. It will be noted that catalysts 8 to 12, each having a silicaalumina support rather than the silica-magnesia support of the first stage catalyst of the present invention, resulted in the desired conversion being obtained only at an unreasonably high average catalyst temperature, in each case 780 F. or above.

TABLE II.-COMPARISON OF FIRST STAGE CATALYST OF PRESENT INVENTION vWITH- CONVENTIONAL CATALYSTS RE ACIDITY, STARTING TEMPERATURE, ISO TO NORMAL C4 PRODUCT RATIO, MIDDLE DIS- TILLATE TO GASOLINE PRODUCT RATIO AND NOR- MAL PARAFFIN CONTENT OF UNCONVERTED BOT- TOMS The following table sets forth on a comparative basis single-stage hydrocraclting results of processing an Arabian straight run feed, at 0.5 LHSV, 2000 p.s.i.a., 60% per-pass conversion to products boiling below the initial boiling point of the feed, and extinction recycle, over the iirst stage catalyst of the present invention, compared with hydrocracking the same feed under the same conditions over various prior art catalysts. The factors compared are: (1) startingl temperature necessary to give said 60% per-pass conversion; (2) the ratio of iC4 to mC., in the product; (3) the ratio Vof 400 to 650 F. product to C5 to 400 F. product, i.e., the ratio of middle distillate production to gasoline production; (4) the hydrogen consumption, in s.c.f. per barrel of feed;

3 0.5 F. per hour.

4 1..0 F. per hour.

cnA

and (5) the change, in F., of the pour point of the smoothly, except in the case of the first stage catalyst of the present invention, with which is obtained a higher ratio than would be expected from inspection of the prior art catalysts alone; (2) as acidity increases, the product middle distillate to gasoline ratio decreases, but remains as high with the first stage catalyst of the present invention as with catalysts of weaker acidity, which is entirely unexpected; heretofore, it has been believed that a catalyst of higher acidity would produce less middle distillate per unit of gasoline production than a more weakly acidic catalyst; (3) as acidity increases, hydrogen consumption increases smoothly, except in the case of the rst stage catalyst of the present invention, with which is obtained a higher hydrogen consumption and improved product quality; (4) as acidity increases, the normal paraffin content of the unconverted bottoms material, as indicated by the F. change in bottoms pour point from the pour point of the feed, decreases and then increases; with Catalysts A and E the bottoms material is indicated to have agreater normal paraffin content than the feed. With Catalysts B, C and D the unconverted bottoms material is less parafiinic than the feed, which is extremely desirable because normal parains vare refractory to hydrocracking and therefore build up in recycle bottoms during recycle operation. A build-up of refractory normal paraffins can effectively prevent the practical use of recycle hydrocracking to produce middle distillates, because prohibitive temperature and pressure increases can be required to crack these refractory compounds; (5) lwith Catalysts B, C and D the undesirable refractory normal paraftins are selectively cracked and/ or are isomei-ized to valuable isoparaffins, to an extent adequate to permit satisfactory recycle operation.

VTABLE III.-COl\IPARISON OF FIRST STAGE CATALYST OF PRESENT INVENTION `WITH CATALYST HAVING SILICA-ALUMINA SUPPORT RE PRODUCTION OF MIDDLE DISTILLATES The following table further indicates the specificity of the first stage catalyst of the present invention for the Start. 40G-650 F./ H2, s.c.f./ Bottoms Cat. Cat. comp. T., F. i iC4/nC4 (J5-400 F. bbl. pour point change, F.

6% Ni+22% Mo on A1203 850 0.2 1. 4 1, 300 +13 NiMo on SiO2-Al20s, 30% S102 765 0. 6 1. 4 1, 700 -38 NiW on SiO2-Mg0, 27% MgO 720 1. l 1. 4 2, 000 -25 NiMO on SiOz-Al2O3, 90% SiO2 790 0. 6 0. 9 1, 800 I-15 6% Ni on Sim-A1203, 90% S1O2 740 1.1 0. 4 2, 600 +19 The catalysts in the above table are set forth in order of increasing acidities, with Catalyst A having the lowest acidity and Catalyst E having the highest acidity. Catalyst C is an example of the iirst'stage catalyst of thepresent invention, while the other catalysts indicated are representative of various prior art catalysts.

production of middle distillates from various hydrocarbon feeds, compared with a catalyst having a silicaalumina support. In this case, the: feed is a 650 to 820 F. hydroined Midway gas oil, containing 3.6 p.p.m. total nitrogen. It is hydrocracked at 0.77 LHSV, 1500 p.s.i.g. and a hydrogen rate of 5000 s.c.f. per barrel of TABLE IV.REGENERABILITY OF FIRST STAGE CATA- LYST OF PRESENT INVENTION AND REGENERATED CATALYST ACTIVITY The following table illustrates the regenerability of the preferred nickel-tungsten on silica-magnesia catalyst of the present invention. A catalyst comprising 7.0% nickel and 19.3% tungsten on a silica-magnesia support containing 27.7% magnesia, with an area of 316 m.2/g., was placed in hydrocracking reactor and contacted for 120 hours at 2000 psig., 1.0 LHSV, 759 F. average catalyst temperature, and hydrogen rate of 5500 s.c.f. per barrel of feed, with a hydrocarbon feed boiling from 650 to 982 F., said feed having a gravity of 23.5 API, an aniline point of 178.9D F., a pour point of +90 ASTM and a total nitrogen content of 665 p.p.m. The catalyst under these conditions converted 54 weight percent of the feed to products boiling below the 650 F. initial boiling point of the feed, and the gravity of the total products produced was 40.3 AFI.

After the foregoing ori-stream period the catalyst was regenerated in a nitrogen-oxygen stream, at a reactor pressure of 600 p.s.i.g. and a gas rate of cubic feet per hour, for a total period of 20 hours. During this period the temperature was slowly raised from 500 to 900 F., and the oxygen content of the gas was raised from 0.1 to 4.0 volume percent.

The regenerated catalyst, having an area of 237 m.2/ g., was then used to hydrocrack the same feed that had been used to hydrocrack prior to regeneration, under the same conditions. The activity of the regenerated catalyst was substantially equal to its original fresh activity, as indicated by its conversion, at an average catalyst temperature of 750 F., of 48 weight percent of the feed to products boiling below the initial boiling point of the feed, the total products produced having a gravity of 38.8 API.

The following summarizes the foregoing results:

TABLE V.-COMPARISON OF FIRST STAGE CATALYST OF PRESENT INVENTION WITH CONVENTIONAL CATALYSTS RE DENITRIFICATION ABILITY, NITRO- GEN SENSITIVITY AND ABILITY TO CONVERT NI- TROGEN-CONTAINING FEEDS TO MIDDLE DISTIL- LATES The following table indicates results obtainable with the first stage catalyst of the present invention and with a low acidity prior art catalyst, and a high acidity prior art catalyst, respectively, when used to hydrocrack a 650 F. to 1000 F. hydrocarbon feed at the indicated temperatures, and at 1.0 LHSV, 2000 p.s.i.g. and a hydrogen rate of 6500 s.c.f. per barrel, with extinction recycle or" unconverted products. The indicated low nitrogenfeeds refer to feeds containing from zero to 10 p.p.m. nitrogen and the indicated high nitrogen feeds refer to feeds containing above 10 ppm. nitrogen, for example 10 to 1000 p.p.m. nitrogen.

NiW on 6% NH- 6% Nl on SiOz-MgO 22% Mo Sim-A1203, 27% Aigo 0H A1203 90% S102 Temperature, in F. for

50% conversion with low feeds 650 850 550 Temperature in F. for

50% conversion with high N feeds 740 850 760 Maximum yield of B20-650 F. middle distillate, with high N feed, percent -85 75-85 55-65 iCt/nCr product ratio high low high Pour point of synthetic middle distillate 1. 3 1. 0 0. 1 nil nil 'nil Sensitivity to N low nil high From the foregoing it will be seen that the process of the present invention is effective to convert a wide range of hydrocarbon feeds to valuable products, mainly middle distillate, light gasoline and isobutane. It will further be seen that, contrary to many prior art processes for producing middle distillates, no heavy gasoline is produced as a final product.

Although only specific embodiments of the present invention have been described, numerous variations could be made in those embodiments without departing from the spirit of the invention, and all such variations that fall within the scope of the appended claims are intended to be embraced thereby.

I claim:

1. In a process for producing gasoline from a hydrocarbon feed selected from the group consisting 0f petroleum distillates boiling from 200 to 1l00 F. and petroleum residua boiling above 500 F., which comprises contacting said feed and from 1000 to 10,000 s.c.f. of hydrogen per barrel of said feed in a first conversion zone at a temperature from 500 to 950 F., a pressure above 500 p.s.i.g. and an LI-ISV of 0.1 to 4.0 With a catalyst comprising at least one hydrogenating component selected from the Group VI metals and compounds of Group VI metals and at least one hydrogenating component selected from the Group VIII metals and compounds of Group VIII metals intimately associated with a catalyst support comprising silica-magnesia, the improvement which comprises recovering from the effluent from said rst conversion zone a light gasoline product boiling below about 180 F., a bottoms product boiling above about 660 F. and a middle distillate product boiling between 300 and 660 F., withdrawing said products from the system, hydrocracking in a second conversion zone in the presence of an active acidic hydrocracldng catalyst at a pressure of 500 to 3000 p.s.i.g. and a temperature of 550 to 850 F. a portion of the effluent from said first conversion zone boiling from about 180 to 300 F., and containing light paraiiins not converted to isoparaihns in the first conversion zone, to produce a light gasoline product having an increased isoparain content, and recovering said light gasoline product from said second conversion zone.

2. A process as in claim 1, wherein said hydrogenating component selected from the Group VI metals and compounds ot Group VI metals is present in an amount from 1 to 40 weight percent, based on the total catalyst composite.

3. A process as in claim 1, wherein said hydrogenating component selected from the Group VIII metals and compounds of Group VIII metals is present in an amount from 1 to 2O weight percent, based on the total catalyst composite.

4. A process as in claim 1 wherein said hydrogenating ,components comprise nickel and tungsten.

5. A process as in claim 1 wherein at least two reactors are used in said first conversion zone, each containing said silicanagnesia supported catalyst, and wherein said reactors are so arranged that they can he switched from parallel, for maximizing middle distillate production, to series, for maximizing gasoline production, whereby the ratio of middle distillate product to gasoline product can be varied.

6. A process as in claim 1 wherein a portion of the eluent from said rst conversion zone that boils above about 660 F. is recycled to that zone.

References Cited by the Examiner UNITED STATES PATENTS 2,428,692 10/47 Voorhies 208-112 Hill 208-95 Johnson et al 208110 Scott et al. 208-80 Ciapetta et al 20E-1110 Haxton et al. 208-65 Folkus 208-59 Mason 20S-112 Myers 20S-59 10 ALPHoNso D. SULLIVAN, Primary Examiner.

UNITED STATES PATENT OFFICE CERTIFICATE 0F CGRRECTION Patent No. 3,180,818 April 27, 1965 William H. Claussen It is hereby certified that error appears in the above numbered patent requiring correction and that the said Letters Patent should read as corrected below.

Column 3, line 26, for "with" read said column 7, line 56, for "or" read on Columns 9 and 10, first table, last column, under the heading "Fouling rate", line 12 thereof,

for ".3" read High.3

Signed and sealed this 28th day of September 1965.

(SEAL) Attest:

ERNEST W. SWIDER EDWARD J. BRENNER Attesting Officer Commissioner of Patents 

1. IN A PROCESS FOR PRODUCING GASOLINE FROM A HYDROCARBON FEED SELECTED FROM THE GROUP CONSISTING OF PETROLEUM DISTILLATES BOILING FROM 200* TO 1100*F. AND PETROLEUM RESIDUA BOILING ABOVE 500*F., WHICH COMPRISES CONTACTING SAID FEED AND FROM 1000 TO 10,000 S.C.F. OF HYDROGEN PER BARREL OF SAID FEED IN A FIRST CONVERSION ZONE AT A TEMPERATURE FROM 500* TO 950*F., A PRESSURE ABOVE 500 P.S.I.G. AND AN LHSV OF 0.1 TO 4.0 WITH A CATALYST COMPRISING AT LEAST ONE HYDROGENATING COMPONENT SELECTED FROM THE GROUP VI METALS AND COMPOUNDS OF GROUP VI METALS AND AT LEAST ONE HYDROGENATING COMPONENT SELECTED FROM THE GROUP VIII METALS AND COMPOUNDS OF GROUP VIII METALS INTIMATELY ASSOCIATED WITH A CATALYST SUPPORT COMPRISING SILICA-MAGNESIA, THE IMPROVEMENT WHICH COMPRISES RECOVERING FROM THE EFFUENT FROM SAID FIRST CONVERSION ZONE A LIGHT GASOLINE PRODUCT BOILING BELOW ABOUT 